Process for Heat Integration for Ethanol Production and Purification Process

ABSTRACT

Ethanol production from the hydrogenation of acetic acid requires energy to drive the hydrogenation reaction and the purification of the crude ethanol product. Heat integration process to recover heat from one part of the production process to be used within the process improves efficiencies and reduces costs.

CROSS REFERENCE TO RELATED APPLICATIONS

This application claims priority to U.S. Provisional App. No.61/300,815, filed on Feb. 2, 2010, U.S. Provisional App. No. 61/332,696,filed on May 7, 2010, and U.S. Provisional App. No. 61/332,699, filed onMay 7, 2010, the entire contents and disclosures of which areincorporated herein by reference.

FIELD OF THE INVENTION

The present invention relates generally to processes for heatintegration in the production of ethanol, and, in particular, heatintegration in the production of ethanol from the hydrogenation ofacetic acid.

BACKGROUND OF THE INVENTION

Ethanol for industrial use is conventionally produced from petrochemicalfeed stocks, such as oil, natural gas, or coal, from feed stockintermediates, such as syngas, or from starchy materials or cellulosematerials, such as corn or sugar cane. Conventional methods forproducing ethanol from petrochemical feed stocks, as well as fromcellulose materials, include the acid-catalyzed hydration of ethylene,methanol homologation, direct alcohol synthesis, and Fischer-Tropschsynthesis. Instability in petrochemical feed stock prices contributes tofluctuations in the cost of conventionally produced ethanol, making theneed for alternative sources of ethanol production all the greater whenfeed stock prices rise. Starchy materials, as well as cellulosematerial, are converted to ethanol by fermentation. However,fermentation is typically used for consumer production of ethanol forfuels or consumption. In addition, fermentation of starchy or cellulosematerials competes with food sources and places restraints on the amountof ethanol that can be produced for industrial use.

Ethanol production via the reduction of alkanoic acids and/or othercarbonyl group-containing compounds has been widely studied, and avariety of combinations of catalysts, supports, and operating conditionshave been mentioned in the literature. During the reduction of alkanoicacid, e.g., acetic acid, other compounds are formed with ethanol or areformed in side reactions. These impurities limit the production andrecovery of ethanol from such reaction mixtures. For example, duringhydrogenation, esters are produced that together with ethanol and/orwater form azeotropes, which are difficult to separate. In addition whenconversion is incomplete, unreacted acid remains in the crude ethanolproduct, which must be removed to recover ethanol.

Conventional methods have been used for separating ethanol fromfermentation processes. These methods have proposed variousthermally-integrated configurations, such as those described in U.S.Pat. Nos. 5,215,902, 5,035,776, 4,626,321, and 4,306,942. However, thesesystems deal with different feed streams and separation requirements.Therefore, a need remains for improving the energy efficiency in therecovery of ethanol from a crude product obtained by reducing alkanoicacids, such as acetic acid, and/or other carbonyl group-containingcompounds.

SUMMARY OF THE INVENTION

In a first embodiment, the present invention is directed to a processfor producing ethanol, the process comprising the steps of introducingan acetic acid vapor feed stream comprising acetic acid into a firstreactor hydrogenating the acetic acid in the first reactor in thepresence of a first catalyst to form a first reactor product streamcomprising ethanol and residual acetic acid. The process furthercomprises cooling the first reactor product stream in a first heatexchange stage, wherein the first reactor product stream is cooled to atemperature that is equal to or greater than a feed temperature to thefirst reactor. In addition, the process comprises hydrogenating theresidual acetic acid in a second reactor in the presence of a secondcatalyst to form a second reactor product stream comprising ethanol; andcooling the second reactor product stream in a second heat exchangestage.

In a second embodiment, the present invention is directed to a processfor producing ethanol, the process comprising the steps of hydrogenatingacetic acid from an acetic acid vapor feed stream in a reactor in thepresence of a catalyst to form a crude reactor product; flashing thecrude reactor product to form a vapor stream and a liquid streamcomprising ethanol, ethyl acetate, water, and acetic acid; andseparating at least a portion of the liquid stream in a firstdistillation column to produce a first distillate comprising ethanol,ethyl acetate and water, and a first residue comprising acetic acid. Theprocess may transfer heat of the first distillate in a first heatexchange stage. The process further comprises separating a condensedportion of the first distillate in a second distillation column toproduce a second distillate comprising ethyl acetate and second residuecomprising ethanol and water; and separating a portion of the secondresidue in a third distillation column to produce a third distillatecomprising ethanol and third residue comprising water. The process maycool a portion of the third residue in a second heat exchange stage.

BRIEF DESCRIPTION OF DRAWINGS

The invention is described in detail below with reference to theappended drawings, wherein like numerals designate similar parts.

FIG. 1 is a schematic diagram of a hydrogenation system in accordancewith one embodiment of the present invention.

FIG. 2 is a thermodynamic model chart of heat integration for ahydrogenation system in accordance with one exemplary embodiment of thepresent invention.

DETAILED DESCRIPTION OF THE INVENTION

The present invention relates to processes for recovering ethanolproduced by hydrogenating acetic acid in the presence of a catalyst.Hydrogenation may also produce one or more byproducts that are separatedin one or more distillation columns. The distillation columns requireheat to separate the byproducts from ethanol. In addition, there areseveral feed streams that require heating before hydrogenation that alsorequire cooling after the reaction. In particular, the present inventionintegrates heat within the process to improve energy efficiency. In oneembodiment, the hydrogenation reactor operates in the vapor phase at atemperature higher than the purification system. Because thehydrogenation reactor produces a crude product in the vapor phase andthe purification is in the liquid phase the heat of the reaction cannotbe readily transferred directly by the crude product. Withouttransferring the heat through the crude product the system may lose heatand reduce the overall efficiency.

One embodiment of the present invention advantageously incorporates heatintegration through one or more heat stages. A heat stage refers tochanging the temperature of a first stream through indirect contact witha second stream, e.g., a cooler first stream is heated by a hottersecond stream through indirect contact of the first and second streams.In some embodiments, a hotter stream may be cooled by transferring heatto a cooler stream. In other embodiments, a cooler stream may be heatedthrough heat transferred by a hotter stream. Heat stages allow theefficient use of the heat produced or used in one part of the system,e.g., the hydrogenation reactor, to be retained within another part ofthe process. The heat stage may use any suitable indirect-contact heatexchangers, including direct transfer type heat exchangers or indirecttransfer type heat exchangers, that are capable of transferring heat byconduction. In some embodiments, a heat transfer fluids, such as steamor oil, may transfer the heat from one stream to another stream. Apreferred heat exchanger that may be used in embodiments of the presentinvention may be a spiral tube heat exchange, double-pipe heatexchanger, shell and tube heat exchanger, or fluidized-bed heatexchanger.

FIG. 1 shows a hydrogenation system 100 suitable for the hydrogenationof acetic acid and separating ethanol from the crude reaction mixtureaccording to one embodiment of the invention. System 100 comprisesreaction zone 101 and distillation zone 102. Acetic acid and hydrogenare fed to a vaporizer 105 via lines 103 and 104, respectively, tocreate a vapor feed stream 106 comprising acetic acid and hydrogen thatis directed to a first reactor 107 and/or second reactor 109.

The raw materials, acetic acid and hydrogen, used in connection with theprocess of this invention may be derived from any suitable sourceincluding natural gas, petroleum, coal, biomass, and so forth. Asexamples, acetic acid may be produced via methanol carbonylation,acetaldehyde oxidation, ethylene oxidation, oxidative fermentation, andanaerobic fermentation. As petroleum and natural gas prices fluctuate,becoming either more or less expensive, methods for producing aceticacid and intermediates such as methanol and carbon monoxide fromalternate carbon sources have drawn increasing interest. In particular,when petroleum is relatively expensive compared to natural gas, it maybecome advantageous to produce acetic acid from synthesis gas (“syngas”) that is derived from any available carbon source. U.S. Pat. No.6,232,352, the disclosure of which is incorporated herein by reference,for example, teaches a method of retrofitting a methanol plant for themanufacture of acetic acid. By retrofitting a methanol plant, the largecapital costs associated with CO generation for a new acetic acid plantare significantly reduced or largely eliminated. All or part of the syngas is diverted from the methanol synthesis loop and supplied to aseparator unit to recover CO and hydrogen, which are then used toproduce acetic acid. In addition to acetic acid, such a process can alsobe used to make hydrogen which may be utilized in connection with thisinvention.

Methanol carbonylation processes suitable for production of acetic acidare described in U.S. Pat. Nos. 7,208,624, 7,115,772, 7,005,541,6,657,078, 6,627,770, 6,143,930, 5,599,976, 5,144,068, 5,026,908,5,001,259, and 4,994,608, the disclosure of which is incorporated hereinby reference. Optionally, the production of ethanol may be heatintegrated with such methanol carbonylation processes.

U.S. Pat. No. RE 35,377, also incorporated herein by reference, providesa method for the production of methanol by conversion of carbonaceousmaterials such as oil, coal, natural gas and biomass materials. Theprocess includes hydrogasification of solid and/or liquid carbonaceousmaterials to obtain a process gas which is steam pyrolized withadditional natural gas to form synthesis gas. The syn gas is convertedto methanol which may be carbonylated to acetic acid. The methodlikewise produces hydrogen which may be used in connection with thisinvention as noted above. U.S. Pat. No. 5,821,111, which discloses aprocess for converting waste biomass through gasification into synthesisgas as well as U.S. Pat. No. 6,685,754, the disclosures of which areincorporated herein by reference.

In one optional embodiment, the acetic acid fed to the hydrogenationreaction may also comprise other carboxylic acids and anhydrides, aswell as acetaldehyde and acetone. Preferably, a suitable acetic acidfeed stream comprises one or more of the compounds selected from thegroup consisting of acetic acid, acetic anhydride, acetaldehyde, ethylacetate, and mixtures thereof. These other compounds may also behydrogenated in the processes of the present invention. In someembodiments, the present of carboxylic acids, such as propanoic acid orits anhydride, may be beneficial in producing propanol.

Alternatively, acetic acid in vapor form may be taken directly as crudeproduct from the flash vessel of a methanol carbonylation unit of theclass described in U.S. Pat. No. 6,657,078, the entirety of which isincorporated herein by reference. The crude vapor product, for example,may be fed directly to the ethanol synthesis reaction zones of thepresent invention without the need for condensing the acetic acid andlight ends or removing water, saving overall processing costs.

In a preferred embodiment, hydrogen in feed line 104 may be preheated bya heat exchanger 146. Acetic acid in feed line 103 may also be preheatedby a heat exchanger 147. Preferably, the heat exchangers 146 and 147 areintegrated with a distillate stream from the distillation zone 102. Thedistillate stream 120 may be from the acetic acid removal column 119that generally operates at a higher temperature than the other columnsto remove heavier components from the crude ethanol product. In the heatexchange stage, the distillate stream 120 may transfer latent heat toacetic acid feed line 103, hydrogen feed line 104, or both, throughindirect heat transfer in heat exchangers 146 and 147, respectively.Preferably distillate stream 120 may be condensed by heat exchangers 146and 147, along with heat exchanger 145. In one embodiment, distillatestream 120 may remain in the vapor phase and may not be condensed untilafter transferring sensible heat to the feed lines 103 and 104. Autility exchanger may further condense distillate stream 120. AlthoughFIG. 1 depicts heat exchangers 146 and 147 in series, heat exchangers146 and 147, along with heat exchanger 145, may preferably be arrangedin parallel.

In other embodiments, heat exchangers 146 and 147 may be heat integratedwith another stream from within the hydrogenation system 101. Anothersuitable heat stage for heat exchangers 146 and/or 147 may involvetransfer of heat from the effluent from one or more of the reactors.Also, a residue stream from one of the distillation columns, such as theacetic acid column 119 and/or product column 127 may also be heatintegrated with heat exchangers 146 and/or 147. Preferably, heatexchangers 146 and 147 are heat integrated with the same stream so thateach stream is preheated to a similar temperature. In one embodiment,acetic acid in feed line 103 may be preheated with the residue in line121 from acetic acid column 119.

In further embodiments, one or more utility exchangers may be used tofurther preheat acetic acid and/or hydrogen in feed lines 103 and 104,respectively. Utility exchangers refer generically to exchangers thatare not integrated with system 100 and receive heat or cooling from anoutside source.

Preferably acetic acid in line 103 may be preheated to a temperature ofat least 40° C., e.g., at least 70° C. or at least 85° C., prior tobeing fed to vaporizer 105. Hydrogen in line 104 preferably is preheatedto a temperature of at least 40° C., e.g., at least 70° C. or at least85° C., prior to being fed to vaporizer 105.

In one embodiment, lines 103 and 104 may be combined and jointly fed tovaporizer 105, e.g., in one stream containing both hydrogen and aceticacid. The acetic acid may be vaporized at the reaction temperature, andthen fed via vapor feed stream 106 to first reactor 107. Vapor feedstream 106 may also comprise hydrogen in an undiluted state or dilutedwith a relatively inert carrier gas, such as nitrogen, argon, helium,carbon dioxide and the like. For reactions run in the vapor phase, thetemperature should be controlled in the system such that it does notfall below the dew point of acetic acid. In one embodiment the aceticacid may be vaporized at the boiling point of acetic acid at theparticular pressure, and then the vaporized acetic acid may be furtherheated to the reactor inlet temperature. In another embodiment, theacetic acid is transferred to the vapor state by passing hydrogen,recycle gas, another suitable gas, or mixtures thereof through theacetic acid at a temperature below the boiling point of acetic acid,thereby humidifying the carrier gas with acetic acid vapors, followed byheating the mixed vapors up to the reactor inlet temperature.Preferably, the acetic acid is transferred to the vapor by passinghydrogen and/or recycle gas through the acetic acid at a temperature ator below 125° C., followed by heating of the combined gaseous stream tothe reactor inlet temperature.

Any feed that is not vaporized is removed from vaporizer 105, as shownin FIG. 1 by blowdown stream 136, and may be recycled thereto.Optionally a portion of stream 136 may be fed to a reboiler that heatsvaporizer 105.

The temperature of vapor feed stream 106 upon exiting vaporizer 105 ispreferably from 100° C. to 250° C., e.g., from 105° C. to 200° C. orfrom 110° C. to 150° C. Preferably, vapor feed stream 106 may bepreheated by one or more heat exchangers before being fed to firstreactor 107. As shown by FIG. 1, vapor feed stream 106 may be preheatedby heat exchangers 140 and/or 141. Heat exchangers 140 and/or 141 areintegrated with the first reactor product 108 from first reactor 107 andthe second reactor product 110 from second reactor 109. The firstreactor product 108 and second reactor product 110 may be genericallyreferred to as the crude reaction product and each has a composition asdescribed below in Table 1. In one embodiment, the first reactor product108 contains relatively more unreacted acetic acid than the secondreactor product 110. It should be understood that some embodiments mayuse additional reactors and the effluent of those additional reactorsmay be heat integrated in a heat exchanger to pre-heat vapor feed stream106. After pre-heating in heat exchangers 140 and 141, vapor feed stream106 may have a temperature of greater than 120° C., e.g., greater than150° C. or greater than 190° C. Preferably, vapor feed stream 106 isheated to a temperature below the reaction temperature in first reactor107.

In preheating vapor feed stream, heat exchangers also cool first reactorproduct 108 and second reactor product 110 of first reactor 107 andsecond reactor 109, respectively. Due to the exothermic hydrogenationreactor, the first reactor product 108 and second reactor product 110exit each reactor at or near the reaction temperature. Preferably, firstreactor product 108 from first reactor 107 is cooled to a temperaturethat is greater than or equal to the feed temperature to first reactor107. The temperature of the first reactor product 108 fed to secondreactor 109 is preferably from 120° C. to 350° C., e.g., from 150° C. to325° C. or from 200° C. to 275° C. Maintaining first reactor product 108above the feed temperature reduces the need for further preheatingbefore being fed to the second reactor 109. The second reactor product110 may be cooled below the feed temperature through one or more heatexchange stages.

Vapor feed stream 106 may be directed to the top of first reactor 107,and first reactor product 108 may be directed to the top of secondreactor 109, as shown in FIG. 1. First reactor 107 and second reactor109 may each comprise one or more reactor beds. There may be heatintegrated between the reactor beds and cooling of the streams betweenthe reactor beds. Lines 106 and 108 may, in some embodiments, bedirected to the side, upper portion, or bottom of either reactor 107 or109. The hydrogenation reaction in the second reactor 109 preferablyuses the unreacted acetic acid and hydrogen in line 108. Although freshreactants may be fed to second reactor 109, it is preferred that secondreactor 109 operates with the unreacted reactants, acetic acid andhydrogen, of the first reactor product 108.

In FIG. 1, prior to be purified or separated, second reactor product 110may be further cooled in one or more heat exchange stages. Preferably,second reactor product 110 may be cooled in heat exchanger 141 withvapor feed stream 106 as discussed above. Because second reactor product110 has a higher temperature than distillation zone 102, it is preferredto also cool second reactor product 110 with one or more heat exchangestages with streams from distillation zone 102. In addition, secondreactor product 110 may also be cooled through a heat exchange with areboiler stream on one of the distillation columns. As shown in FIG. 1,it is preferred to transfer heat to reboiler stream 126 of light endscolumn 123 through heat exchanger 142. In addition, as shown in FIG. 1,second reactor product 110 may be cooled in heat exchanger 143 withrecycled compressed vapor stream 114. The order of the heat exchangestages for cooling second reactor product 110 may vary depending on thelayout of the facility. In other embodiments, second reactor product 110may be cooled by heating another stream, such as the acetic acid and/orhydrogen feed lines or the liquid feed stream to distillation zone 102.

In first and second reactors 107 and 109 the hydrogenation of aceticacid forms ethanol and water in equal molar ratio, along with one ormore byproducts. Suitable hydrogenation catalysts include catalystscomprising a first metal and optionally one or more metals, on acatalyst support. The first and optional metals may be selected fromGroup IB, IIB, IIIB, IVB, VB, VIIB, VIIB, VIII transitional metals, alanthanide metal, an actinide metal or a metal selected from any ofGroups IIIA, IVA, VA, and VIA. Preferred metal combinations for someexemplary catalyst compositions include platinum/tin,platinum/ruthenium, platinum/rhenium, palladium/ruthenium,palladium/rhenium, cobalt/palladium, cobalt/platinum, cobalt/chromium,cobalt/ruthenium, silver/palladium, copper/palladium, nickel/palladium,gold/palladium, ruthenium/rhenium, and ruthenium/iron. Exemplarycatalysts are further described in U.S. Pat. No. 7,608,744 and U.S. Pub.Nos. 2010/0029995 and 2010/0197485, the entireties of which areincorporated herein by reference.

In one exemplary embodiment, the catalyst comprises a first metalselected from the group consisting of copper, iron, cobalt, nickel,ruthenium, rhodium, palladium, osmium, iridium, platinum, titanium,zinc, chromium, rhenium, molybdenum, and tungsten. Preferably, the firstmetal is selected from the group consisting of platinum, palladium,cobalt, nickel, and ruthenium. More preferably, the first metal isselected from platinum and palladium. When the first metal comprisesplatinum, it is preferred that the catalyst comprises platinum in anamount less than 5 wt. %, e.g., less than 3 wt. % or less than 1 wt. %,due to the high demand for platinum.

As indicated above, the catalyst optionally further comprises a secondmetal, which typically would function as a promoter. If present, thesecond metal preferably is selected from the group consisting of copper,molybdenum, tin, chromium, iron, cobalt, vanadium, tungsten, palladium,platinum, lanthanum, cerium, manganese, ruthenium, rhenium, gold, andnickel. More preferably, the second metal is selected from the groupconsisting of copper, tin, cobalt, rhenium, and nickel. More preferably,the second metal is selected from tin and rhenium.

If the catalyst includes two or more metals, e.g., a first metal and asecond metal, the first metal optionally is present in the catalyst inan amount from 0.1 to 10 wt. %, e.g., from 0.1 to 5 wt. %, or from 0.1to 3 wt. %. The second metal preferably is present in an amount from 0.1and 20 wt. %, e.g., from 0.1 to 10 wt. %, or from 0.1 to 5 wt. %. Forcatalysts comprising two or more metals, the two or more metals may bealloyed with one another or may comprise a non-alloyed metal solution ormixture.

The preferred metal ratios may vary depending on the metals used in thecatalyst. In some exemplary embodiments, the mole ratio of the firstmetal to the second metal is from 10:1 to 1:10, e.g., from 4:1 to 1:4,from 2:1 to 1:2, from 1.5:1 to 1:1.5 or from 1.1:1 to 1:1.1.

The catalyst may also comprise a third metal selected from any of themetals listed above in connection with the first or second metal, solong as the third metal is different from the first and second metals.In preferred aspects, the third metal is selected from the groupconsisting of cobalt, palladium, ruthenium, copper, zinc, platinum, tin,and rhenium. More preferably, the third metal is selected from cobalt,palladium, and ruthenium. When present, the total weight of the thirdmetal preferably is from 0.05 and 4 wt. %, e.g., from 0.1 to 3 wt. %, orfrom 0.1 to 2 wt. %.

In addition to one or more metals, the exemplary catalysts furthercomprise a support or a modified support, meaning a support thatincludes a support material and a support modifier, which adjusts theacidity of the support material. The total weight of the support ormodified support, based on the total weight of the catalyst, preferablyis from 75 wt. % to 99.9 wt. %, e.g., from 78 wt. % to 97 wt. %, or from80 wt. % to 95 wt. %. In preferred embodiments that use a modifiedsupport, the support modifier is present in an amount from 0.1 wt. % to50 wt. %, e.g., from 0.2 wt. % to 25 wt. %, from 0.5 wt. % to 15 wt. %,or from 1 wt. % to 8 wt. %, based on the total weight of the catalyst.

Suitable support materials may include, for example, stable metaloxide-based supports or ceramic-based supports. Preferred supportsinclude silicaceous supports, such as silica, silica/alumina, a GroupIIA silicate such as calcium metasilicate, pyrogenic silica, high puritysilica, and mixtures thereof. Other supports may include, but are notlimited to, iron oxide, alumina, titania, zirconia, magnesium oxide,carbon, graphite, high surface area graphitized carbon, activatedcarbons, and mixtures thereof.

In the production of ethanol, the catalyst support may be modified witha support modifier that is either acidic or basic. Preferably, thesupport modifier is a modifier that has a low volatility or novolatility. Suitable basic modifiers, for example, may be selected fromthe group consisting of: (i) alkaline earth oxides, (ii) alkali metaloxides, (iii) alkaline earth metal metasilicates, (iv) alkali metalmetasilicates, (v) Group JIB metal oxides, (vi) Group IIB metalmetasilicates, (vii) Group IIIB metal oxides, (viii) Group IIIB metalmetasilicates, and mixtures thereof. In addition to oxides andmetasilicates, other types of modifiers including nitrates, nitrites,acetates, and lactates may be used. Preferably, the support modifier isselected from the group consisting of oxides and metasilicates of any ofsodium, potassium, magnesium, calcium, scandium, yttrium, and zinc, aswell as mixtures of any of the foregoing. Preferably, the basic supportmodifier is a calcium silicate, and more preferably calcium metasilicate(CaSiO₃). If the support modifier comprises calcium metasilicate, it ispreferred that at least a portion of the calcium metasilicate is incrystalline form.

Suitable acidic modifiers include those selected from the groupconsisting of TiO₂, ZrO₂, Nb₂O₅, Ta₂O₅, Al₂O₃, B₂O₃, P₂O₅, Sb₂O₃, WO₃,MoO₃, Fe₂O₃, Cr₂O₃, V₂O₅, MnO₂, CuO, Co₂O₃, and Bi₂O₃.

A preferred silica support material is SS61138 High Surface Area (HSA)Silica Catalyst Carrier from Saint-Gobain N or Pro. The Saint-Gobain Nor Pro SS61138 silica contains approximately 95 wt. % high surface areasilica; a surface area of about 250 m²/g; a median pore diameter ofabout 12 nm; an average pore volume of about 1.0 cm³/g as measured bymercury intrusion porosimetry and a packing density of about 0.352 g/cm³(22 lb/ft³).

A preferred silica/alumina support material is KA-160 (Sud Chemie)silica spheres having a nominal diameter of about 5 mm, a density ofabout 0.562 g/ml, in absorptivity of about 0.583 g H₂O/g support, asurface area of about 160 to 175 m²/g, and a pore volume of about 0.68ml/g.

As will be appreciated by those of ordinary skill in the art, supportmaterials are selected such that the catalyst system is suitably active,selective and robust under the process conditions employed for theformation of ethanol.

The metals of the catalysts may be dispersed throughout the support,coated on the outer surface of the support (egg shell) or decorated onthe surface of the support.

The catalyst compositions suitable for use with the present inventionpreferably are formed through metal impregnation of the modifiedsupport, although other processes such as chemical vapor deposition mayalso be employed. Such impregnation techniques are described in U.S.Pat. No. 7,608,744, U.S. Pub. Nos. 2010/0029995 and 2010/0197485,referred to above, the entireties of which are incorporated herein byreference.

In general the hydrogenation reactors, e.g., first reactor 107 andsecond reactor 109, may include a variety of configurations using afixed bed reactor vessel or a fluidized bed reactor vessel, as one ofskill in the art will readily appreciate. Preferably at least tworeactor vessels are used. Each reactor vessel may have one or morereactor beds. In one embodiment each reactor vessel may have at leasttwo reactor beds. Preferably at least four reactor beds are used inembodiments of the present invention to achieve high conversion ofacetic acid. In many embodiments of the present invention, an“adiabatic” reactor can be used; that is, there is little or no need forinternal plumbing through the reaction zone to add or remove heat. Inother embodiments, radial flow reactor or reactors may be employed, or aseries of reactors may be employed with or with out heat exchange,quenching, or introduction of additional feed material. Alternatively, ashell and tube reactor provided with a heat transfer medium may be used.

In preferred embodiments, the catalyst is employed in a fixed bedreactor, e.g., in the shape of a pipe or tube, where the reactants,typically in the vapor form, are passed over or through the catalyst.Other reactors, such as fluid or ebullient bed reactors, can beemployed. In some instances, the hydrogenation catalysts may be used inconjunction with an inert material to regulate the pressure drop of thereactant stream through the catalyst bed and the contact time of thereactant compounds with the catalyst particles.

The hydrogenation reaction may be carried out in either the liquid phaseor vapor phase. Preferably, the reaction is carried out in the vaporphase under the following conditions. The reaction temperature may rangefrom 125° C. to 350° C., e.g., from 200° C. to 325° C., from 225° C. to300° C., or from 250° C. to 300° C. The pressure may range from 10 KPato 3000 KPa (about 1.5 to 435 psi), e.g., from 50 KPa to 2300 KPa, orfrom 100 KPa to 1500 KPa. The reactants may be fed to the reactor at agas hourly space velocity (GHSV) of greater than 500 hr⁻¹, e.g., greaterthan 1000 hr⁻¹, greater than 2500 hr⁻¹ or even greater than 5000 hr⁻¹.In terms of ranges the GHSV may range from 50 hr⁻¹ to 50,000 hr⁻¹, e.g.,from 500 hr⁻¹ to 30,000 hr⁻¹, from 1000 hr⁻¹ to 10,000 hr⁻¹, or from1000 hr⁻¹ to 6500 hr⁻¹. When the reaction is carried out in multiplereactors, each reactor preferably operates at similar conditions.

The hydrogenation optionally is carried out at a pressure justsufficient to overcome the pressure drop across the catalytic bed at theGHSV selected, although there is no bar to the use of higher pressures,it being understood that considerable pressure drop through the reactorbed may be experienced at high space velocities, e.g., 5000 hr⁻¹ or 6500hr⁻¹.

Although the reaction consumes two moles of hydrogen per mole of aceticacid to produce one mole of ethanol, the actual molar ratio of hydrogento acetic acid in the feed stream may vary from about 100:1 to 1:100,e.g., from 50:1 to 1:50, from 20:1 to 1:2, or from 12:1 to 1:1. Mostpreferably, the molar ratio of hydrogen to acetic acid is greater than2:1, e.g., greater than 4:1 or greater than 8:1.

Contact or residence time can also vary widely, depending upon suchvariables as amount of acetic acid, catalyst, reactor, temperature andpressure. Typical contact times range from a fraction of a second tomore than several hours when a catalyst system other than a fixed bed isused, with preferred contact times, at least for vapor phase reactions,of from 0.1 to 100 seconds, e.g., from 0.3 to 80 seconds or from 0.4 to30 seconds.

In particular, the hydrogenation of acetic acid may achieve favorableconversion of acetic acid and favorable selectivity and productivity toethanol. For purposes of the present invention, the term “conversion”refers to the amount of acetic acid in the feed that is converted to acompound other than acetic acid. Conversion is expressed as a molepercentage based on acetic acid in the feed. The conversion may be atleast 10%, e.g., at least 20%, at least 40%, at least 50%, at least 60%,at least 70% or at least 80%. Although catalysts that have highconversions are desirable, such as at least 80% or at least 90%, in someembodiments a low conversion may be acceptable at high selectivity forethanol. It is, of course, well understood that in many cases, it ispossible to compensate for conversion by appropriate recycle streams oruse of larger reactors, but it is more difficult to compensate for poorselectivity.

Selectivity is expressed as a mole percent based on converted aceticacid. It should be understood that each compound converted from aceticacid has an independent selectivity and that selectivity is independentfrom conversion. For example, if 50 mole % of the converted acetic acidis converted to ethanol, we refer to the ethanol selectivity as 50%.Preferably, the catalyst selectivity to ethoxylates is at least 60%,e.g., at least 70%, or at least 80%. As used herein, the term“ethoxylates” refers specifically to the compounds ethanol,acetaldehyde, and ethyl acetate. Preferably, the selectivity to ethanolis at least 80%, e.g., at least 85% or at least 88%. Preferredembodiments of the hydrogenation process also have low selectivity toundesirable products, such as methane, ethane, and carbon dioxide. Theselectivity to these undesirable products preferably is less than 4%,e.g., less than 2% or less than 1%. More preferably, these undesirableproducts are not detectable. Formation of alkanes may be low, andideally less than 2%, less than 1%, or less than 0.5% of the acetic acidpassed over the catalyst is converted to alkanes, which have littlevalue other than as fuel.

The term “productivity,” as used herein, refers to the grams of aspecified product, e.g., ethanol, formed during the hydrogenation basedon the kilograms of catalyst used per hour. A productivity of at least200 grams of ethanol per kilogram catalyst per hour, e.g., at least 400or at least 600. In terms of ranges, the productivity preferably is from200 to 3,000 grams of ethanol per kilogram catalyst per hour, e.g., from400 to 2,500 or from 600 to 2,000.

In various embodiments, the crude ethanol product produced by thehydrogenation process, before any subsequent processing, such aspurification and separation, will typically comprise unreacted aceticacid, ethanol and water. As used herein, the term “crude ethanolproduct” generally refers to any composition comprising from 5 to 70 wt.% ethanol and from 5 to 35 wt. % water. Exemplary embodiments of crudeethanol compositional ranges are provided in Table 1.

TABLE 1 CRUDE ETHANOL PRODUCT COMPOSITIONS Conc. Conc. Conc. Conc.Component (wt. %) (wt. %) (wt. %) (wt. %) Ethanol 5 to 70 10 to 60  15to 50 25 to 50 Acetic Acid 0 to 90 5 to 80 15 to 70 20 to 70 Water 5 to35 5 to 30 10 to 30 10 to 26 Ethyl Acetate 0 to 20 0 to 15  1 to 12  3to 10 Acetaldehyde 0 to 10 0 to 3  0.1 to 3   0.2 to 2   Others 0.1 to10   0.1 to 6   0.1 to 4   —

In addition, multiple reactor vessels and/or reactor beds may be usedthat each preferably contain a similar hydrogenation catalyst. In oneembodiment, a similar catalyst may be used in each of the reactors 107and 109. Preferably, in each reactor the catalyst comprises similarmetals, but the amounts of the metals on the catalyst may vary betweenthe reactor vessels and/or reactor beds. In addition, the supports mayalso vary. In one embodiment, one or more guard beds (not shown) may beused to protect the catalyst from poisons or undesirable impuritiescontained in the feed or return/recycle streams. Such guard beds may beemployed in the vapor or liquid streams. Suitable guard bed materialsare known in the art and include, for example, carbon, silica, alumina,ceramic, or resins. In one aspect, the guard bed media is functionalizedto trap particular species such as sulfur or halogens. During thehydrogenation process, a crude ethanol product is withdrawn, preferablycontinuously, from reactor 107 via line 108 and from reactor 109 vialine 110. The crude ethanol product in line 110, after cooling throughheat exchangers 141, 142 and 143, may be further condensed and fed to ahigh pressure flasher 111 that provides a vapor stream and a liquidstream. Flasher 111 in one embodiment preferably operates at atemperature of from 50° C. to 500° C., e.g., from 70° C. to 400° C. orfrom 100° C. to 350° C. In one embodiment, the pressure of flasher 111preferably is from 50 KPa to 2000 KPa, e.g., from 75 KPa to 1500 KPa orfrom 100 to 1000 KPa.

Vapor stream 112 exiting flasher 111 may comprise hydrogen andhydrocarbons, which may be purged and/or returned to reaction zone 101.As shown in FIG. 1, the returned portion of vapor stream 112 passesthrough compressor 113. Compressed stream 114 may be preheated by heatexchanger 143 and fed to vaporizer 105. In optional embodiments, thecompressed stream 114 may be preheated by heat exchanger 146 along withfresh hydrogen from feed line 104.

The liquid from flasher 111 is withdrawn via liquid stream 115 and fedto a low pressure flasher 116. Low pressure flasher 116 preferablyoperates at a pressure of from 0.1 KPa to 1000 KPa, e.g., from 0.1 KPato 500 KPa or from 0.1 KPa to 100 KPa. In one embodiment, the pressureof flasher 116 preferably is at least 50 KPa lower than flasher 111,e.g., at least 100 KPa lower or at least 200 KPa lower. Flasher 116 mayalso operate at a lower temperature and/or pressure than flasher 111. Inone embodiment, the temperature of flasher 116 preferably is from 20° C.to 100° C., e.g., from 30° C. to 85° C. or from 40° C. to 70° C. In oneembodiment, the temperature of flasher 116 preferably is at least 50° C.lower than flasher 111, e.g., at least 75° C. lower or at least 100° C.lower.

The vapor stream 117 exiting flasher 116 may comprise hydrogen andhydrocarbons, and is preferably purged. Optionally, vapor stream 117 maybe returned to the reaction zone 101 in a manner similar to that ofvapor stream 112. The liquid feed stream 118 from flasher 116 iswithdrawn and fed to first column 119.

In one embodiment, liquid feed stream 118 may be preheated in one ormore heat exchange stages before being separated in first column 119.Preferably, liquid feed stream 118 is heated to a temperature from 50°C. to 180° C., e.g., 80° C. to 140° C. or from 100° C. to 120° C. Liquidfeed stream 118 may be preheated in two heat exchange stages by heatexchangers 144 and 145. In heat exchanger 144 liquid feed stream 118 isindirectly heated by the third residue 129 from the third column 127,referred to as the product column. In heat exchanger 145 liquid feedstream 118 is indirectly heated by distillate stream 120 of first column119. Preferably distillate stream 120 may be condensed by heat exchanger145. In one embodiment, distillate stream 120 may remain in the vaporphase and may not be condensed until after transferring the heat toliquid feed stream 118. Although FIG. 1 depicts heat exchangers 145,146, and 147 in series, heat exchangers 145, 146, and 147 may preferablybe placed in parallel.

In some embodiments, low pressure flasher 116 may be bypassed, in partor in whole, and liquid stream 115 may be fed to distillation zone 102.When liquid stream 115 is fed to distillation zone, it is preferred toheat liquid stream in a manner similar to liquid feed stream 118, asdescribed above.

The contents of stream 118, and stream 115, typically may besubstantially similar to the product obtained directly from the reactor,and may, in fact, also be characterized as a crude ethanol product.However, liquid feed stream 118 preferably has substantially nohydrogen, carbon dioxide, methane or ethane, which are removed byflashers 111 and/or 116. Exemplary components of liquid feed stream 118,as well as liquid stream 115, are provided in Table 2. It should beunderstood that stream 115 and/or 118 may contain other components, notlisted, such as components in the feed.

TABLE 2 FEED COMPOSITION Conc. (wt. %) Conc. (wt. %) Conc. (wt. %)Ethanol 5 to 70    10 to 60 15 to 50 Acetic Acid <90     5 to 80 15 to70 Water 5 to 35    5 to 30 10 to 30 Ethyl Acetate <20   0.001 to 15  1to 12 Acetaldehyde <10  0.001 to 3 0.1 to 3   Acetal <5 0.001 to 2 0.005to 1    Acetone <5  0.0005 to 0.05 0.001 to 0.03  Other Esters <5 <0.005<0.001 Other Ethers <5 <0.005 <0.001 Other Alcohols <5 <0.005 <0.001

The amounts indicated as less than (<) in the tables throughout presentapplication are preferably not present and if present may be present intrace amounts or in amounts greater than 0.0001 wt. %.

The “other esters” in Table 2 may include, but are not limited to, ethylpropionate, methyl acetate, isopropyl acetate, n-propyl acetate, n-butylacetate or mixtures thereof. The “other ethers” in Table 2 may include,but are not limited to, diethyl ether, methyl ethyl ether, isobutylethyl ether or mixtures thereof. The “other alcohols” in Table 2 mayinclude, but are not limited to, methanol, isopropanol, n-propanol,n-butanol or mixtures thereof. In one embodiment, the feed composition,e.g., line 118, may comprise propanol, e.g., isopropanol and/orn-propanol, in an amount from 0.001 to 0.1 wt. %, from 0.001 to 0.05 wt.% or from 0.001 to 0.03 wt. %. In should be understood that these othercomponents may be carried through in any of the distillate or residuestreams described herein and will not be further described herein,unless indicated otherwise.

When the content of acetic acid in line 118 is less than 5 wt. %, theacid separation column 119 may be skipped and line 118 may be introduceddirectly to second column 123, also referred to herein as a light endscolumn.

In the embodiment shown in FIG. 1, line 118 is introduced in the lowerpart of first column 119, e.g., lower half or lower third. In firstcolumn 119, unreacted acetic acid, a portion of the water, and otherheavy components, if present, are removed from the composition in line118 and are withdrawn, preferably continuously, as residue. Some or allof the residue may be returned and/or recycled back to reaction zone 101via line 121. A portion of the residue may be fed to a reboiler (notshown) via stream 121′ to provide heat to column 119. The reboiler ofcolumn 119 may integrated with a suitable stream, such as first reactorproduct 108 and/or second reactor product 110.

First column 119 also forms an overhead distillate, which is withdrawnin line 120. Prior to condensing distillate stream 120, distillate maybe directed to one or more heat exchange stages. The one or more heatstages may comprise heat exchangers arranged in parallel. As discussedabove, it is preferred that distillate may be used to heat liquid feedstream 118 fed to column 119, acetic acid feed line 103 and hydrogenfeed line 104. During these heat exchange stages, distillate stream 120may be condensed to the liquid phase and refluxed to column 119, forexample, at a ratio of from 10:1 to 1:10, e.g., from 3:1 to 1:3 or from1:2 to 2:1.

Any of columns 119, 123, 127, or optional column 132 in distillationzone 102 may comprise any distillation column capable of separationand/or purification. The columns preferably comprise tray columns havingfrom 1 to 150 trays, e.g., from 10 to 100 trays, from 20 to 95 trays orfrom 30 to 75 trays. The trays may be sieve trays, fixed valve trays,movable valve trays, or any other suitable design known in the art. Inother embodiments, a packed column may be used. For packed columns,structured packing or random packing may be employed. The trays orpacking may be arranged in one continuous column or they may be arrangedin two or more columns such that the vapor from the first section entersthe second section while the liquid from the second section enters thefirst section, etc.

The associated condensers and liquid separation vessels that may beemployed with each of the distillation columns may be of anyconventional design and are simplified in FIG. 1. As shown in FIG. 1,heat may be supplied to the base of each column or to a circulatingbottom stream through a heat exchanger or reboiler. Other types ofreboilers, such as internal reboilers, may also be used in someembodiments. The heat that is provided to reboilers may be derived fromany heat generated during the process that is integrated with thereboilers or from an external source such as another heat generatingchemical process or a boiler. In addition to the reactors and flashersshown in FIG. 1, additional reactors, flashers, condensers, heatingelements, and other components may be used in embodiments of the presentinvention. As will be recognized by those skilled in the art, variouscondensers, pumps, compressors, reboilers, drums, valves, connectors,separation vessels, etc., normally employed in carrying out chemicalprocesses may also be combined and employed in the processes of thepresent invention.

The temperatures and pressures employed in any of the columns may vary.As a practical matter, pressures from 10 KPa to 3000 KPa will generallybe employed in these zones although in some embodiments subatmosphericpressures may be employed as well as superatmospheric pressures.Temperatures within the various zones will normally range between theboiling points of the composition removed as the distillate and thecomposition removed as the residue. It will be recognized by thoseskilled in the art that the temperature at a given location in anoperating distillation column is dependent on the composition of thematerial at that location and the pressure of column. In addition, feedrates may vary depending on the size of the production process and, ifdescribed, may be generically referred to in terms of feed weightratios.

When first column 119 is operated under standard atmospheric pressure,the temperature of the first residue stream 121 preferably is from 95°C. to 120° C., e.g., from 105° C. to 117° C. or from 110° C. to 115° C.The temperature of the distillate stream 120 exiting from column 119preferably is from 70° C. to 110° C., e.g., from 75° C. to 95° C. orfrom 80° C. to 90° C. In other embodiments, the pressure of first column119 may range from 0.1 KPa to 510 KPa, e.g., from 1 KPa to 475 KPa orfrom 1 KPa to 375 KPa. Exemplary components of the distillate andresidue compositions for first column 119 are provided in Table 3 below.It should also be understood that the distillate and residue may alsocontain other components, not listed, such as components in the feed.For convenience, the distillate and residue of the first column may alsobe referred to as the “first distillate” or “first residue.” Thedistillates or residues of the other columns may also be referred towith similar numeric modifiers (second, third, etc.) in order todistinguish them from one another, but such modifiers should not beconstrued as requiring any particular separation order.

TABLE 3 FIRST COLUMN Conc. (wt. %) Conc. (wt. %) Conc. (wt. %)Distillate Ethanol 20 to 75 30 to 70 40 to 65 Water 10 to 40 15 to 35 20to 35 Acetic Acid <2 0.001 to 0.5  0.01 to 0.2  Ethyl Acetate <60 5.0 to40  10 to 30 Acetaldehyde <10 0.001 to 5    0.01 to 4   Acetal <0.1 <0.1<0.05 Acetone <0.05 0.001 to 0.03   0.01 to 0.025 Residue Acetic Acid 60 to 100 70 to 95 85 to 92 Water <30  1 to 20  1 to 15 Ethanol <1 <0.9<0.07

Depending on the reaction conditions, the crude ethanol product exitingreactors 107 and/or 109 in line 112 may comprise ethanol, acetic acid(unconverted), ethyl acetate, and water. After exiting reactors 107and/or 109, a non-catalyzed equilibrium reaction may occur between thecomponents contained in the crude ethanol product until it is added toflasher 111 and/or first column 119. This equilibrium reaction tends todrive the crude ethanol product to an equilibrium between ethanol/aceticacid and ethyl acetate/water.

In the event the crude ethanol product is temporarily stored, e.g., in aholding tank, prior to being directed to distillation zone 102, extendedresidence times may be encountered. Additional heat integration may beused to cool the crude ethanol product when stored in the holding tankand heat when withdrawn from the holding tank. Generally, the longer theresidence time between reaction zone 101 and distillation zone 102, thegreater the formation of ethyl acetate. For example, when the residencetime between reaction zone 101 and distillation zone 102 is greater than5 days, significantly more ethyl acetate may form at the expense ofethanol. Thus, shorter residence times between reaction zone 101 anddistillation zone 102 are generally preferred in order to maximize theamount of ethanol formed. In one embodiment, a holding tank (not shown),is included between reaction zone 101 and distillation zone 102 fortemporarily storing the liquid component from stream 115 and/or stream118 for up to 5 days, e.g., up to 1 day, or up to 1 hour. In a preferredembodiment no tank is included and the condensed liquids are feddirectly to the first distillation column 119. In addition, the rate atwhich the non-catalyzed reaction occurs may increase as the temperatureof the crude ethanol product, e.g., in streams 115 and/or 118,increases. These reaction rates may be particularly problematic attemperatures exceeding 30° C., e.g., exceeding 40° C. or exceeding 50°C. Thus, in one embodiment, the temperature of liquid components instream 115 or in the optional holding tank is maintained at atemperature less than 40° C., e.g., less than 30° C. or less than 20° C.One or more cooling devices may be used to reduce the temperature of theliquid in streams 115 and/or 118.

As discussed above, a holding tank (not shown) may be included betweenthe reaction zone 101 and distillation zone 102 for temporarily storingsome or all of the streams 115 and/or 118, for example from 1 to 24hours, optionally at a temperature of about 21° C., and corresponding toan ethyl acetate formation of from 0.01 wt. % to 1.0 wt. % respectively.In addition, the rate at which the non-catalyzed reaction occurs mayincrease as the temperature of the crude ethanol product is increased.For example, as the temperature of the crude ethanol product in streams115 and/or 118 increases from 4° C. to 21° C., the rate of ethyl acetateformation may increase from about 0.01 wt. % per hour to about 0.005 wt.% per hour. Thus, in one embodiment, the temperature of liquidcomponents in streams 115 and/or 118 or in the optional holding tank ismaintained at a temperature less than 21° C., e.g., less than 4° C. orless than −10° C. In one embodiment, after liquid in streams 115 and/or118 is stored, it may be preheated in one or more heat exchange stages.

First distillate stream 120 preferably comprises ethanol, ethyl acetate,and water, along with other impurities, which may be difficult toseparate due to the formation of binary and tertiary azeotropes. Thecondensed first distillate 122 is introduced to the second column 123,also referred to as the “light ends column,” preferably in the middlepart of column 123, e.g., middle half or middle third. As one example,when a 25 tray column is utilized in a column without water extraction,condensed first distillate 122 is introduced at tray 17. In oneembodiment, the second column 123 may be an extractive distillationcolumn. In such embodiments, an extraction agent, such as water, may beadded to second column 123. An extraction agent may also be obtainedfrom an external source. Preferably, third reside 129 which compriseswater from the third column 127 may be used as the extractive agent.Prior to feeding the extraction agent to the second column 123, thirdresidue 129 may be cooled in one or more heat exchange stages.Preferably, third residue 129 is cooled by indirectly heating liquidfeed stream 118 in heat exchanger 144. Third residue 129 may also becooled by indirectly heating a reboiler. Third residue 129 may be cooledto a temperature of from 35° C. to 100° C., e.g., 65° C. to 95° C. orfrom 75° C. to 85° C. Without cooling third residue 129, third residue129 would tend to vaporize the feed to second column 123 and thuslimiting extractive distillation.

Second column 123 may be a tray column or packed column. In oneembodiment, second column 123 is a tray column having from 5 to 70trays, e.g., from 15 to 50 trays or from 20 to 45 trays. Although thetemperature and pressure of second column 123 may vary, when atatmospheric pressure the temperature of the second residue 124preferably is from 60° C. to 90° C., e.g., from 70° C. to 90° C. or from80° C. to 90° C. The temperature of the second distillate 125 preferablyis from 50° C. to 90° C., e.g., from 60° C. to 80° C. or from 60° C. to70° C. Column 123 may operate at atmospheric pressure. In otherembodiments, the pressure of second column 123 may range from 0.1 KPa to510 KPa, e.g., from 1 KPa to 475 KPa or from 1 KPa to 375 KPa. Exemplarycomponents for the distillate and residue compositions for second column123 are provided in Table 4 below. It should be understood that thedistillate and residue may also contain other components, not listed,such as components in the feed.

TABLE 4 SECOND COLUMN Conc. (wt. %) Conc. (wt. %) Conc. (wt. %)Distillate Ethyl Acetate 10 to 90 25 to 90 50 to 90 Acetaldehyde  1 to25  1 to 15 1 to 8 Water  1 to 25  1 to 20  4 to 16 Ethanol <30 0.001 to15   0.01 to 5   Acetal <5 0.001 to 2    0.01 to 1   Residue Water 30 to70 30 to 60 30 to 50 Ethanol 20 to 75 30 to 70 40 to 70 Ethyl Acetate <30.001 to 2    0.001 to 0.5  Acetic Acid <0.5 0.001 to 0.3  0.001 to 0.2 

The weight ratio of ethanol in the second residue to ethanol in thesecond distillate preferably is at least 3:1, e.g., at least 6:1, atleast 8:1, at least 10:1 or at least 15:1. The weight ratio of ethylacetate in the second residue to ethyl acetate in the second distillatepreferably is less than 0.4:1, e.g., less than 0.2:1 or less than 0.1:1.In embodiments that use an extractive column with water as an extractionagent as the second column 123, the weight ratio of ethyl acetate in thesecond residue to ethyl acetate in the second distillate approacheszero.

A portion of the residue in line 125 may be heated in reboiler stream126 through heat exchanger 142 to supply the energy for column 123. Asdiscussed above, heat exchanger 142 preferably is integrated with thesecond reactor product 110. Optionally, heat exchanger 142 may also beheat integrated with other streams, such as the residue of anothercolumn, and in particular column 119 and/or column 127. Additionalutility exchangers (not shown) may also be used to control start upconditions.

As shown, the second residue from the bottom of second column 123, whichcomprises ethanol and water, is fed via line 125 to third column 127,also referred to as the “product column.” More preferably, the secondresidue in line 125 is introduced in the lower part of third column 127,e.g., lower half or lower third. Third column 127 recovers ethanol,which preferably is substantially pure other than the azeotropic watercontent, as the distillate stream 128. The distillate of third column127 preferably is refluxed as shown in FIG. 1, for example, at a refluxratio of from 1:10 to 10:1, e.g., from 1:3 to 3:1 or from 1:2 to 2:1.The third residue stream 129, which preferably comprises primarilywater, preferably is removed from the system 100 or may be partiallyreturned to any portion of the system 100. Third column 127 ispreferably a tray column as described above and preferably operates atatmospheric pressure. The temperature of the third distillate exiting inline 128 from third column 127 preferably is from 60° C. to 110° C.,e.g., from 70° C. to 100° C. or from 75° C. to 95° C. The temperature ofthe third residue 129 exiting from third column 127 preferably is from70° C. to 115° C., e.g., from 80° C. to 110° C. or from 85° C. to 105°C., when the column is operated at atmospheric pressure. Exemplarycomponents of the distillate and residue compositions for third column127 are provided in Table 5 below. It should be understood that thedistillate and residue may also contain other components, not listed,such as components in the feed.

TABLE 5 THIRD COLUMN Conc. (wt. %) Conc. (wt. %) Conc. (wt. %)Distillate Ethanol 75 to 96  80 to 96 85 to 96 Water <12 1 to 9 3 to 8Acetic Acid <1 0.001 to 0.1  0.005 to 0.01  Ethyl Acetate <5 0.001 to4    0.01 to 3   Residue Water 75 to 100  80 to 100  90 to 100 Ethanol<0.8 0.001 to 0.5  0.005 to 0.05  Ethyl Acetate <1 0.001 to 0.5  0.005to 0.2  Acetic Acid <2 0.001 to 0.5  0.005 to 0.2 

Any of the compounds that are carried through the distillation processfrom the feed or crude reaction product generally remain in the thirddistillate in amounts of less 0.1 wt. %, based on the total weight ofthe third distillate composition, e.g., less than 0.05 wt. % or lessthan 0.02 wt. %. In one embodiment, one or more side streams may removeimpurities from any of the columns 119, 123 and/or 127 in the system100. Preferably at least one side stream is used to remove impuritiesfrom the third column 127. The impurities may be purged and/or retainedwithin the system 100.

The third residue 129 may be used as an extractive agent for column 123.When used as an extractive agent, it may be preferred to cool the thirdresidue 129 in one or more heat exchange stages. Preferably, asdiscussed above, third residue 129 is cooled indirectly heating liquidfeed stream 118 in heat exchanger 144. In addition, a portion of thirdresidue 129 may be directed to a reboiler via stream 130. This portionin stream 130 may be heated in one or more heat exchange stages. Also, aportion of the third residue 129 may be purged from the system 100 viastream 131. Depending on the subsequent use of purge stream 131, thepurge stream 131 may be cooled.

In optional embodiments, the third distillate in line 128 may be cooledby in a heat exchange stage with one or more cool streams. Thirddistillate 128 may optionally be cooled indirectly heating liquid feedstream 118 in a heat exchanger (not shown). Third distillate may also becooled indirectly heating the acetic acid feed line 103 and/or hydrogenfeed line 104 to vaporizer 105 in a heat exchanger (not shown).

The third distillate in line 128 may be further purified to form ananhydrous ethanol product stream, i.e., “finished anhydrous ethanol,”using one or more additional separation systems, such as, for example,distillation columns (e.g., a finishing column) or molecular sieves.

In optional embodiments, distillate 121 of first column 119 may be heatintegrated with a reboiler (not shown) of third column 127. When usingthis heat integration, first column 119 may operate at an elevatedpressure.

Returning to second column 123, the second distillate preferably isrefluxed as shown in FIG. 1, for example, at a reflux ratio of from 1:10to 10:1, e.g., from 1:5 to 5:1 or from 1:3 to 3:1. In one embodimentsecond distillate 124 may be fed to reaction zone 101, and in particularvaporizer 105 and/or first reactor 107. Optional purges may be takenfrom second distillate 124 to remove build up of impurities.

In optional embodiments, the second distillate is fed via line 124 tooptional fourth column 132, also referred to as the “acetaldehyderemoval column.” In fourth column 132 the second distillate is separatedinto a fourth distillate, which comprises acetaldehyde, in line 133 anda fourth residue, which comprises ethyl acetate, in line 134. The fourthdistillate 133 preferably is refluxed at a reflux ratio of from 1:20 to20:1, e.g., from 1:15 to 15:1 or from 1:10 to 10:1, and a portion offourth distillate 133 may be returned to the reaction zone 101. Forexample, fourth distillate 133 may be combined with acetic acid feedline 103, added to vaporizer 105, or added directly to first reactor107. Without being bound by theory, since acetaldehyde may behydrogenated to form ethanol, the recycling of a stream that containsacetaldehyde to the reaction zone increases the yield of ethanol anddecreases byproduct and waste generation. In another embodiment (notshown in the figure), the acetaldehyde may be collected and used, withor without further purification, to make useful products including butnot limited to n-butanol, 1,3-butanediol, and/or crotonaldehyde andderivatives.

The fourth residue of fourth column 132 may be purged via line 134. Aportion of fourth residue in line 134 may be directed to a reboilerstream 135. The fourth residue primarily comprises ethyl acetate andethanol, which may be suitable for use as a solvent mixture or in theproduction of esters. Fourth residue 134 may be hydrolyzed and thehydrolyzed products may be directed within the system 100. In onepreferred embodiment, the acetaldehyde is removed from the seconddistillate in fourth column 134 such that no detectable amount ofacetaldehyde is present in the residue of column 134.

Optional fourth column 132 is preferably a tray column as describedabove and preferably operates above atmospheric pressure. In oneembodiment, the pressure is from 120 KPa to 5,000 KPa, e.g., from 200KPa to 4,500 KPa, or from 400 KPa to 3,000 KPa. In a preferredembodiment the fourth column 132 may operate at a pressure that ishigher than the pressure of the other columns.

The temperature of the fourth distillate exiting in line 134 from fourthcolumn 132 preferably is from 60° C. to 110° C., e.g., from 70° C. to100° C. or from 75° C. to 95° C. The temperature of the residue exitingfrom fourth column 132 preferably is from 70° C. to 115° C., e.g., from80° C. to 110° C. or from 85° C. to 110° C. Exemplary components of thedistillate and residue compositions for fourth column 132 are providedin Table 6 below. It should be understood that the distillate andresidue may also contain other components, not listed, such ascomponents in the feed.

TABLE 6 FOURTH COLUMN Conc. (wt. %) Conc. (wt. %) Conc. (wt. %)Distillate Acetaldehyde 2 to 80  2 to 50 5 to 40 Ethyl Acetate <90 30 to80 40 to 75  Ethanol <30 0.001 to 25   0.01 to 20   Water <25 0.001 to20   0.01 to 15   Residue Ethyl Acetate 40 to 100  50 to 100 60 to 100Ethanol <40 0.001 to 30   0 to 15 Water <25 0.001 to 20   2 to 15Acetaldehyde  <1 0.001 to 0.5  Not detectable Acetal  <3 0.001 to 2   0.01 to 1   

The finished ethanol composition obtained by the processes of thepresent invention preferably comprises from 75 to 96 wt. % ethanol,e.g., from 80 to 96 wt. % or from 85 to 96 wt. % ethanol, based on thetotal weight of the finished ethanol composition. Exemplary finishedethanol compositional ranges are provided below in Table 7.

TABLE 7 FINISHED ETHANOL COMPOSITIONS Component Conc. (wt. %) Conc. (wt.%) Conc. (wt. %) Ethanol 75 to 96 80 to 96 85 to 96 Water <12 1 to 9 3to 8 Acetic Acid <1 <0.1 <0.01 Ethyl Acetate <2 <0.5 <0.05 Acetal <0.05<0.01 <0.005 Acetone <0.05 <0.01 <0.005 Isopropanol <0.5 <0.1 <0.05n-propanol <0.5 <0.1 <0.05

The finished ethanol composition of the present invention preferablycontains very low amounts, e.g., less than 0.5 wt. %, of other alcohols,such as methanol, butanol, isobutanol, isoamyl alcohol and other C₄-C₂₀alcohols. In one embodiment, the amount of isopropanol in the finishedethanol is from 80 to 1,000 wppm, e.g., from 95 to 1,000 wppm, from 100to 700 wppm, or from 150 to 500 wppm. In one embodiment, the finishedethanol composition preferably is substantially free of acetaldehyde andmay comprise less than 8 wppm of acetaldehyde, e.g., less than 5 wppm orless than 1 wppm.

The finished ethanol composition produced by the embodiments of thepresent invention may be used in a variety of applications includingfuels, solvents, chemical feedstocks, pharmaceutical products,cleansers, sanitizers, hydrogenation transport or consumption. In fuelapplications, the finished ethanol composition may be blended withgasoline for motor vehicles such as automobiles, boats and small pistonengine aircrafts. In non-fuel applications, the finished ethanolcomposition may be used as a solvent for toiletry and cosmeticpreparations, detergents, disinfectants, coatings, inks, andpharmaceuticals. The finished ethanol composition may also be used as aprocessing solvent in manufacturing processes for medicinal products,food preparations, dyes, photochemicals and latex processing.

The finished ethanol composition may also be used a chemical feedstockto make other chemicals such as vinegar, ethyl acrylate, ethyl acetate,ethylene, glycol ethers, ethylamines, aldehydes, and higher alcohols,especially butanol. In the production of ethyl acetate, the finishedethanol composition may be esterified with acetic acid or reacted withpolyvinyl acetate. The finished ethanol composition may be dehydrated toproduce ethylene. Any of known dehydration catalysts can be employed into dehydrate ethanol, such as those described in U.S. Pub. Nos.2010/0030002 and 2010/0030001, the entire contents and disclosures ofwhich are hereby incorporated by reference. A zeolite catalyst, forexample, may be employed as the dehydration catalyst. Preferably, thezeolite has a pore diameter of at least about 0.6 nm, and preferredzeolites include dehydration catalysts selected from the groupconsisting of mordenites, ZSM-5, a zeolite X and a zeolite Y. Zeolite Xis described, for example, in U.S. Pat. No. 2,882,244 and zeolite Y inU.S. Pat. No. 3,130,007, the entireties of which are hereby incorporatedby reference.

In order that the invention disclosed herein may be more efficientlyunderstood, an example is provided below. The following example describethe various heat integration processes of the present invention.

EXAMPLE

A hydrogenation system, as constructed in FIG. 1, was thermodynamicallymodeled in FIG. 2 to integrate heat within the system. Each arrowrepresents a stream in the system that is either heated or cooled. Inthe thermodynamic model, there are five zones, 151-155, in which heatmay be heat integrated between any of the hotter streams with the coolerstreams.

In zone 151, two hot streams, first reactor product 108 and secondreactor product 110, are cooled. Also there are several cool streams,vapor feed stream 106, reboiler stream 121′, compressed vapor stream114, acetic acid feed line 103, hydrogen feed line 104, reboiler stream135, reboiler stream 130, and blowdown stream 136. First reactor product108 exits first reactor 107 at a temperature of 292.3° C. and is cooledby heat exchanger 140 to a temperature of 250° C. Second reactor product110 exits the second reactor 109 at a temperature of 279.5° C. and iscooled by heat exchanger 141 to a temperature of 141.4° C. Secondreactor product 110 is subsequently cooled in zones 152 and 153. Heatexchangers 140 and 141 in turn heats vapor feed stream 106, from aninitial temperature of 119.6° C. to 250° C.

As indicated by zone 151, additional heat integration of first reactorproduct 108 and second reactor product 110 with the other streams may beemployed depending on the operating conditions of the system. Inparticular, the utility exchangers may be replaced by one or more heatexchange stages. For example, first reactor product 108 and/or secondreactor product 110 may be heat integrated with reboiler streams 121′,135, and/or 130, as well as with blowdown stream 136.

In zone 152, there are two hot streams, second reactor product 110 andthird residue 129. Also there are several cool streams, compressed vaporstream 114, acetic acid feed line 103, hydrogen feed line 104, reboilerstream 126, and liquid feed stream 118. Second reactor product 110 fromzone 151 has a temperature of 141.4° C. and is cooled by heat exchanger142 to a temperature of 115.6° C. Heat exchangers 142 heats reboilerstream 126 to temperature of about 103.8° C. Subsequently, secondreactor product 110 may be further cooled by heat exchanger 143 to atemperature of 102.8° C. Heat exchanger 143 heats compressed vaporstream 114 from a temperature of 48.6° C. to 104.4° C. Compressed vaporstream 114 may further be heated by a utility exchanger in zone 151.Third residue may also be cooled to a temperature of 80° C. by heatexchanger 144. Heat exchanger 144 heats liquid feed stream 118 from atemperature of 67.8° C. to 110° C. Liquid feed stream 118 is alsopreheated in zone 153 by a heat exchanger 145.

In zone 153, in addition to hot streams second reactor product 110 andthird residue 129 from zone 151, there are several other hot streamsfrom the distillate streams of the columns. These hot streams includethe first distillate 120, third distillate 128, second distillate 124,and optional fourth distillate 133. In addition, hot streams in zone 153comprise vapor streams 112 and 117. The cool streams comprise compressedvapor stream 114, acetic acid feed line 103, hydrogen feed line 104, andliquid feed stream 118. First distillate 120 is used to transfer heatthrough heat exchangers 145, 146, and 147. First distillate 120 exitsfirst column 119 at a temperature of about 96° C. may be transfer heatin heat exchangers 145, 146, and 147. First distillate 120 may also besubsequently condensed. Heat exchanger 145 heats liquid feed stream 118from a temperature of 53.3° C. to 67.8° C. and liquid is subsequentlyheated by heat exchanger 144. Heat exchangers 146 and 147 heats hydrogenfeed line 104 and acetic acid feed line 103, respectively, from atemperature of 32.2° C. to 85° C. Hydrogen feed line 104 and acetic acidfeed line 103 are further heated by one or more utility exchangers inzones 151 and 152 prior to be introduced to vaporizer 105.

Although not used in this example, the other distillate streams may alsobe heat integrated in a similar manner as first distillate stream 120.

In zones 154 and 155 there are no cool streams shown in FIG. 2. However,these zones may be candidates for heat integration when colder hotstreams are present. This may include intermediate storage wheretemperature may be very low. It is also a possible to heat integratewith scrubbers that require using a large amount of process water at alow temperature.

While the invention has been described in detail, modifications withinthe spirit and scope of the invention will be readily apparent to thoseof skill in the art. In view of the foregoing discussion, relevantknowledge in the art and references discussed above in connection withthe Background and Detailed Description, the disclosures of which areall incorporated herein by reference. In addition, it should beunderstood that aspects of the invention and portions of variousembodiments and various features recited below and/or in the appendedclaims may be combined or interchanged either in whole or in part. Inthe foregoing descriptions of the various embodiments, those embodimentswhich refer to another embodiment may be appropriately combined withother embodiments as will be appreciated by one of skill in the art.Furthermore, those of ordinary skill in the art will appreciate that theforegoing description is by way of example only, and is not intended tolimit the invention.

1-30. (canceled)
 31. A process for producing ethanol, the processcomprising the steps of: introducing an acetic acid vapor feed streamcomprising acetic acid into a reactor; hydrogenating the acetic acid inthe reactor in the presence of a catalyst to form a reactor productstream; flashing the reactor product stream to form a vapor stream and aliquid stream comprising ethanol, ethyl acetate, water, and acetic acid;transferring at least part of the heat from the reactor product streamin a first heat exchange stage to the acetic acid vapor feed stream;transferring at least part of the heat from the reactor product streamin a second heat exchange stage to the vapor feed stream; and recoveringethanol from the liquid stream.
 32. The process of claim 31, furthercomprising separating at least a part of the liquid stream in one ormore distillation columns to recover ethanol.
 33. The process of claim32, further comprising transferring at least part of the heat from thereactor product stream in a third heat exchange stage to the liquidstream prior to being introduced to the one or more distillationcolumns.
 34. The process of claim 32, further comprising transferring atleast part of the heat from the reactor product stream in a fourth heatexchange stage to a reboiler stream of at least one of the one or moredistillation columns.
 35. The process of claim 31, further comprisingintroducing an acetic acid feed stream, a hydrogen feed stream, and thevapor stream to a vaporizer to produce the acetic acid vapor feedstream.
 36. The process of claim 35, further comprising transferring atleast part of the heat from the reactor product stream in a third heatexchange stage to the acetic acid feed stream.
 37. The process of claim35, further comprising transferring at least part of the heat from thereactor product stream in a third heat exchange stage to the hydrogenfeed stream.
 38. The process of claim 31, wherein the reactor productstream is cooled to a temperature that is below a feed temperature tothe reactor.
 39. The process of claim 31, wherein the feed temperatureis from 100° C. to 250° C.
 40. The process of claim 31, wherein thefirst heat exchange stage comprises one or more indirect-contact heatexchangers.
 41. The process of claim 31, wherein the second heatexchange stage comprises one or more indirect-contact heat exchangers.42. The process of claim 31, wherein the reactor comprises one or morereactor beds.
 43. A process for producing ethanol, the processcomprising the steps of: introducing an acetic acid feed stream and ahydrogen feed stream to a vaporizer to produce an acetic acid vapor feedstream; hydrogenating acetic acid from the acetic acid vapor feed streamin a reactor in the presence of a catalyst to form a crude reactorproduct; flashing the crude reactor product to form a vapor stream and aliquid stream comprising ethanol, ethyl acetate, water, and acetic acid;separating at least a part of the liquid stream in a distillation columnto yield a distillate and a residue; transferring at least part of theheat of the distillate in a first heat exchange stage to the liquidstream; transferring at least part of the heat of the distillate in asecond heat exchange stage to the acetic acid feed stream; andrecovering ethanol from the distillation column.
 44. The process ofclaim 43, further comprising transferring at least part of the heat ofthe distillate in a third heat exchange stage to the hydrogen feedstream.
 45. The process of claim 43, wherein the distillate is notcondensed after the first heat exchange stage and second heat exchangestage.
 46. The process of claim 43, wherein the distillate comprisesethanol, ethyl acetate, and water.
 47. The process of claim 43, furthercomprising introducing the vapor stream to the vaporizer to produce theacetic acid vapor feed stream.
 48. A process for producing ethanol, theprocess comprising the steps of: introducing an acetic acid feed streamand a hydrogen feed stream to a vaporizer to produce an acetic acidvapor feed stream; hydrogenating acetic acid from the acetic acid vaporfeed stream in a reactor in the presence of a catalyst to form a crudereactor product; flashing the crude reactor product to form a vaporstream and a liquid stream comprising ethanol, ethyl acetate, water, andacetic acid; separating at least a part of the liquid stream in adistillation column to yield a distillate and a residue; transferring atleast part of the heat of the distillate in a first heat exchange stageto the liquid stream; transferring at least part of the heat of thedistillate in a second heat exchange stage to the hydrogen feed stream;and recovering ethanol from the distillation column.
 49. The process ofclaim 48, further comprising transferring at least part of the heat ofthe distillate in a third heat exchange stage to the acetic acid feedstream.
 50. The process of claim 48, wherein the distillate is notcondensed after the first heat exchange stage and second heat exchangestage.